Dehydrogenation of dehydrogenatable hydrocarbons

ABSTRACT

A process for the catalytic dehydrogenation of a dehydrogenatable C 2  -plus feed hydrocarbon which comprises the steps of: (a) passing a feed stream comprising the C 2  -plus feed hydrocarbon into a dehydrogenation zone and through at least one bed of dehydrogenation catalyst maintained at dehydrogenation conditions and producing a dehydrogenation zone effluent stream comprising hydrogen, the C 2  -plus feed hydrocarbon and a C 2  -plus product hydrocarbon; (b) forming an oxidation catalyst bed feed stream by admixing an oxygen-containing stream with the dehydrogenation zone effluent stream; (c) passing the oxidation catalyst bed feed stream through a bed of hydrogen selective oxidation catalyst maintained at selective oxidation conditions and producing an oxidation zone effluent stream having a reduced concentration of hydrogen; and (d) recovering the product hydrocarbon from the oxidation zone effluent stream without contacting the oxidation zone effluent stream with dehydrogenation catalyst.

BACKGROUND OF THE INVENTION

The field of art to which this invention pertains is the catalyticdehydrogenation of a dehydrogenatable hydrocarbon to produce an olefinicor aromatic hydrocarbon. The preferred use of the subject method is inthe dehydrogenation of alkylaromatic hydrocarbons such as the conversionof ethylbenzene to styrene. The invention is specifically related to theinjection of an oxygen-containing gas into a bed of selective hydrogenoxidation catalyst located in the dehydrogenation zone effluent stream.

INFORMATION DISCLOSURE

The dehydrogenation of hydrocarbons is well described in the prior art,with both acyclic and aromatic hydrocarbons being thereby converted tothe corresponding less saturated products. For instance, dehydrogenationis performed commercially for the production of styrene fromethylbenzene to fulfill the sizable demand for this polymer precursor.U.S. Pat. No. 3,515,766 (Root et al) and U.S. Pat. No. 3,409,689 (Ward)show typical prior art catalytic steam dehydrogenation processes foralkylaromatics including ethylbenzene. These references describe theadmixture of superheated steam into the feed hydrocarbon and theadmixture of additional amounts of superheated steam with the reactantsbetween sequential beds of dehydrogenation catalyst to reheat thereactants.

The dehydrogenation of low molecular weight paraffin hydrocarbons is ahighly developed process. For instance, processes for thedehydrogenation of paraffins are described in U.S. Pat. Nos. 3,391,218;3,448,165; 3,649,566; 3,647,911; and 3,714,281. These referencesdescribe various catalysts and process conditions which may be employed.

A typical prior art process flow comprises the admixture of the feedhydrocarbon with hydrogen and the heating of the feed stream throughindirect heat exchange with the dehydrogenation zone effluent stream.The feed stream may comprise recycled unconverted hydrocarbons and willnormally comprise recycled hydrogen. After being heated in thefeed-effluent heat exchanger, the feed stream is further heated bypassage through a heater which is typically a fired heater or furnace.The feed stream is then contacted with a bed of dehydrogenationcatalyst, which may be either a fixed, moving or fluidized bed ofcatalyst. The dehydrogenation reaction is very endothermic and theentering reactants are quickly cooled to temperatures at which thedehydrogenation reaction does not proceed at an acceptable rate. Tocounteract this cooling effect of the reaction, heat may be supplied tothe bed of dehydrogenation catalyst by indirect heat exchange withcirculating high temperature fluids or by a rapid turnover of catalystin a fluidized bed system.

Another method of supplying the necessary heat of reaction is to removethe reactants from the bed of dehydrogenation catalyst and to heat thereactants externally through the use of a heater. In this instance thereactants which emerge from the first bed of dehydrogenation catalystare passed through a heater which may be similar to the initial feedheater. The thus-heated reactants are then passed through a second bedof dehydrogenation catalyst. This contacting-reheating sequence may berepeated as many times as desired. A process for the dehydrogenation ofethylbenzene utilizing interstage reheating of the reactants isdescribed in U.S. Pat. No. 2,959,626.

Still another method of reheating the reactants in a multistagedehydrogenation process is through the use of superheated steam, whichcan be admixed into the feed stream to the first reaction stage and/orto each subsequent reaction stage. This type of interstage reheating isnormally associated with the dehydrogenation of alkylaromatichydrocarbons and is described in U.S. Pat. No. 3,515,766.

Whatever form the reaction zone takes, it is customary for the effluentstream of the dehydrogenation reaction zone to be passed through thefeed-effluent heat exchanger for heat recovery and to then be cooledsufficiently to cause a partial condensation of the effluent stream. Thepartial condensation facilitates the easy separation of the bulk of thehydrogen from the other components of the effluent stream, with aportion of the hydrogen being removed as a net product gas and a secondportion normally being recycled to the dehydrogenation reaction zone.The remaining mixture of saturated and unsaturated hydrocarbons andby-products is bypassed into the appropriate products recoveryfacilities, which will typically comprise a first stripping column whichremoves light ends having boiling points below that of the desiredproduct and a second fractionation column which separates the remaininghydrocarbons into product and recycle streams.

It is also known in the prior art to pass oxygen into a dehydrogenationzone for the purpose of reacting the oxygen with hydrogen releasedduring the dehydrogenation reaction to thereby liberate heat and toconsume hydrogen. The processes known to employ this technique utilize ahydrogen oxidation catalyst in an attempt to selectively oxidize thehydrogen rather than feed or product hydrocarbons also present in thedehydrogenation zone. For instance, U.S. Pat. No. 3,437,703 (Reitmeir etal) discloses a dehydrogenation process which may utilize either a"homogeneous catalyst system" in which oxidation and dehydrogenationcatalysts are admixed or a layered system of individual catalyst bedsreferred to as a "multi-space bed" system.

Two other references also disclose the utilization of oxygen within adehydrogenation zone. U.S. Pat. No. 3,502,737 (Ghublikian) presents aprocess for the dehydrogenation of ethylbenzene which indicates catalystactivity and stability are maintained by the careful control of theamount of oxygen which is present and by reduction in the steam which isused in the reaction zone. An oxygen-containing gas such as air issupplied both initially and at interstage points in a carefullycontrolled manner. U.S. Pat. No. 3,855,330 (Mendelsohn et al) alsodiscloses a dehydrogenation process using sequential beds ofdehydrogenation catalyst and oxidation catalyst. According to theteachings of the '330 reference, it is preferred that oxygen isintroduced only after substantial conversion of the feed hydrocarbon,that it is desirable that oxygen does not come into contact with thedehydrogenation catalyst, and that the major part or all of the addedoxygen is consumed within the bed of oxidation catalyst.

The use of multi-stage reaction systems in which the catalyst movesdownward by gravity flow between different reaction stages is disclosedin U.S. Pat. Nos. 3,761,390 and 3,907,511. These systems are basicallydirected to the reforming of petroleum naphthas, but may be adapted tothe processing of other hydrocarbons.

In U.S. Pat. No. 4,418,237 (Imai), a process is disclosed for thedehydrogenation of a dehydrogenatable hydrocarbon which comprisescontacting the dehydrogenatable hydrocarbon with a dehydrogenationcatalyst at dehydrogenation conditions in the presence of steam,contacting the resulting mixture of undehydrogenated dehydrogenatablehydrocarbon, resultant dehydrogenated hydrocarbon, hydrogen and steam anoxygen-containing gas in the presence of an oxidation catlyst atoxidation conditions to selectively oxidize hydrogen and recovering saiddehydrogenated hydrocarbon.

U.S. Pat. No. 4,435,607 (Imai), a process is disclosed for thedehydrogenation of a dehydrogenatable hydrocarbon with separate andintermediate selective oxidation of hydrogen to thereby raise thetemperature of the unconverted and dehydrogenated hydrocarbons beforeintroduction into a subsequent dehydrogenation zone containingdehydrogenation catalyst.

It is believed that there has heretofore been no attempt or descriptionof oxidizing at least a portion of the resulting hydrogen contained in adehydrogenation zone effluent stream by contacting the effluent streamin admixture with an oxygen-containing stream with a hydrogen selectiveoxidation catalyst and providing an oxidation zone effluent streamhaving a reduced concentration of hydrogen and recovering the producthydrocarbon from the oxidation zone effluent stream without contactingthe oxidation zone effluent stream with dehydrogenation catalyst. Thecited references appear silent in this respect.

BRIEF SUMMARY OF THE INVENTION

The subject invention provides a means of minimizing the heat input andthe volume of gas to be compressed by the net gas compressor and whichinvention employs selective hydrogen combustion of the dehydrogenationzone effluent stream. The invention increases the amount of heat whichis available to be exchanged to heat the fresh feed hydrocarbon streamand reduces the volume of gas which must be compressed thereby loweringthe cost of utilities and improving the overall economy of the processof the present invention.

One broad embodiment of the invention may be characterized as a processfor the catalytic dehydrogenation of a dehydrogenatable C₂ -plus feedhydrocarbon which comprises the steps of: (a) passing a feed streamcomprising the C₂ -plus feed hydrocarbon into a dehydrogenation zone andthrough at least one bed of dehydrogenation catalyst maintained atdehydrogenation conditions and producing a dehydrogenation zone effluentstream comprising hydrogen, the C₂ -plus feed hydrocarbon and a C₂ -plusproduct hydrocarbon; (b) forming an oxidation catalyst bed feed streamby admixing an oxygen-containing stream with the dehydrogenation zoneeffluent stream; (c) passing the oxidation catalyst bed feed streamthrough a bed of hydrogen selective oxidation catalyst maintained atselective oxidation conditions and producing an oxidation zone effluentstream having a reduced concentration of hydrogen; and (d) recoveringthe product hydrocarbon from the oxidation zone effluent stream withoutcontacting the oxidation zone effluent stream with dehydrogenationcatalyst.

Another embodiment of the invention may be characterized as a processfor the catalytic dehydrogenation of a dehydrogenatable C₂ -plus feedhydrocarbon which comprises the steps of: (a) passing a feed streamcomprising the C₂ -plus feed hydrocarbon into a dehydrogenation zone andthrough at least one bed of dehydrogenation catalyst maintained atdehydrogenation conditions and producing a dehydrogenation zone effluentstream comprising hydrogen, the C₂ -plus feed hydrocarbon and a C₂ -plusproduct hydrocarbon; (b) forming an oxidation catalyst bed feed streamby admixing an oxygen-containing stream with the dehydrogenation zoneeffluent stream; (c) passing the oxidation catalyst bed feed streamthrough a bed of hydrogen selective oxidation catalyst maintained atselective oxidation conditions and producing an oxidation zone effluentstream having a reduced concentration of hydrogen; (d) heat exchangingthe oxidation zone effluent stream with the feed stream in step (a) toraise the temperature of the feed stream and to lower the temperature ofthe oxidation zone effluent stream; and (e) recovering the producthydrocarbon from the oxidation zone effluent stream from step (d)without contacting the oxidation zone effluent stream withdehydrogenation catalyst.

Other embodiments of the subject invention encompass further detailssuch as particular hydrocarbonaceous charge stocks, dehydrogenationcatalysts, selective hydrogen oxidation catalysts and operatingconditions, all of which are hereinafter disclosed in the followingdiscussion of each of these facets of the invention.

BRIEF DESCRIPTION OF THE DRAWING

The drawing is a simplified process flow diagram of a preferredembodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The dehydrogenation of low molecular weight hydrocarbons to produce thecorresponding unsaturated hydrocarbons is practiced commercially toprovide feedstocks for the petroleum and petrochemical industries. It isexpected that there will be increased utilization of these processes inthe near future to supply the olefins required for the production ofincreased amounts of lead-free high octane motor fuels, and asfeedstocks for various petrochemical operations.

Processes for the dehydrogenation of aromatic hydrocarbons are also inwidespread commercial use. For instance, large quantities of styrene areproduced by the dehydrogenation of ethylbenzene. The resultant styrenemay be polymerized with itself or it may be copolymerized withbutadiene, isoprene, acrylonitrile, etc. Other hydrocarbons which may bedehydrogenated in much the same manner include diethylbenzene, ethyltoluene, propyl benzene, and isopropyl benzene.

The present process can also be applied to the dehydrogenation of othertypes of hydrocarbons including relatively pure or mixed streams of C₂-C₁₆ paraffins. The process can therefore be applied to thedehydrogenation of propane, butanes, hexanes or nonanes. However, sincethe great majority of the present commercial dehydrogenation processesare employed for the dehydrogenation of ethylbenzene, the followingdescription of the subject invention will be presented primarily interms of the dehydrogenation of ethylbenzene. This is not intended toexclude from the scope of the subject invention those alkylaromatic andacyclic hydrocarbons set out above or those having different ringstructures including bicyclic compounds.

Regardless of the type of feedstock charged to a dehydrogenation processand the type or design of the actual catalytic dehydrogenation zone, theresulting effluent from the catalytic dehydrogenation zone compriseshydrogen, unconverted feed hydrocarbon and dehydrogenated hydrocarbonproduct. This resulting effluent must then be subjected to cooling andfractionation to provide for product separation and purification. Duringthe product separation, the net gas, including primarily the producedhydrogen, must necessarily be compressed for removal from thedehydrogenation process and for subsequent use. In accordance with thepresent invention, the resulting effluent from the catalyticdehydrogenation zone is admixed with an oxygen-containing stream andthen introduced into a selective oxidation zone containing oxidationcatalyst whereby at least a portion of the hydrogen is selectivelyoxidized thereby reducing the volume of net gas which must be compressedand providing additional heat which may then be recovered and utilizedin the process. The present invention thereby provides an improveddehydrogenation process having lower overall utility consumption andtherefore more economical operation.

As described hereinabove, an oxygen-containing gas stream is admixedwith the effluent of the dehydrogenation zone and the resultingadmixture is passed into a bed of selective hydrogen oxidation catalystcontained in a selective oxidation zone. To achieve the optimum level ofperformance and safety in the present process, it is necessary toclosely control the rate at which oxygen is passed into the process inthis manner. An insufficient amount of oxygen will result in a less thandesired consumption of hydrogen and a less than desired reheating of theflowing stream. It is not normally desired to inject an excess amount ofoxygen above that required to perform the desired degree of hydrogencombustion. The passage of an excess amount of oxygen into the selectiveoxidation zone will have detrimental effects upon the long termoperation of the process. Operation of the selective oxidation zone in amanner which does not result in the total consumption of the oxygen isundesirable because of the obvious explosive nature ofoxygen-hydrocarbon mixtures. The explosive nature of these mixtures can,however, be essentially negated by properly operating the process toavoid the presence of mixtures being within the explosive range, asthrough the use of diluents and intentionally low oxygen addition rates,and the presence of a sufficient amount of solid material to act as anexplosion suppression means. Lastly, the presence of oxygen is notnormally desired in vessels containing hydrocarbons as the oxygen mayreact with the hydrocarbons to form various undesired oxygenatedhydrocarbonaceous compounds.

In the drawing, one embodiment of the present invention is illustratedby means of a simplified flow diagram in which such details as pumps,instrumentation, heat exchange and heat-recovery circuits, compressorsand similar hardware have been deleted as being non-essential to theunderstanding of the techniques involved. The use of such miscellaneousappurtenances and recycle streams are well within the purview of oneskilled in the art of petroleum refining techniques. With reference nowto the drawing, a feed stream comprising relatively high purityethylbenzene carried by conduit 1 is introduced into the process andheated in feed/effluent heat exchanger 8 and is subsequently admixedwith superheated steam from conduit 2 and passed into thedehydrogenation zone 3 via conduit 1. At least a portion of the feedadmixture is dehydrogenated to produce styrene and hydrogen.Dehydrogenation zone 3 may be comprised of a multiplicity ofdehydrogenation stages. Regardless of the details associated withdehydrogenation zone 3, the effluent from dehydrogenation zone 3 istransported via conduit 4 and is admixed with a high-purity oxygenstream which is introduced via conduit 5. This resulting mixture isintroduced via conduit 4 into selective oxidation zone 6 which containsa catalyst which promotes the selective combustion or oxidation of thehydrogen released in the dehydrogenation zone to thereby consume thehydrogen and release heat. By controlling the amount of oxygen which isadded through conduit 5, the extent to which the total amount ofavailable hydrogen is combusted within selective oxidation zone 6 mayalso be controlled. This control is preferably performed on the basis ofa temperature measurement taken at the outlet of the selective oxidationzone 6. The rate of oxygen addition through conduit 5 is thereforepreferably controlled on the basis of the preferred inlet temperature tofeed/effluent heat exchanger 8. The effluent from selective oxidationzone 6 is removed via conduit 7 and introduced into feed/effluent heatexchanger 8 wherein the feed ethylbenzene stream is heated as describedhereinabove. The cooled stream from feed/effluent heat exchanger 8 istransported via conduit 7 and is introduced into fractionation zone 13.A low pressure net gas stream comprising normally gaseous hydrocarbonsand unoxidized hydrogen, if any, is removed from fractionation zone 13via conduit 9, compressed in net gas compressor 10 and recovered viaconduit 9. A normally liquid hydrocarbonaceous stream comprisingunconverted feed hydrocarbon, ethylbenzene, and dehydrogenated feedhydrocarbon, styrene, is removed from fractionation zone 13 via conduit11 and is introduced into a conventional styrene recovery facility. Anaqueous stream comprising condensed feed steam and the water resultingfrom the selective oxidation of hydrogen is removed from fractionationzone 13 via conduit 12 and recovered.

The subject invention has the advantage of reducing the quantity ofnoncondensible gas which must be compressed and removed from theprocess. A second advantage is the reduction of utility consumptionwhich is realized by the lower quantity of utilities required tocompress the net gas and the ability to recover the heat regenerated bythe selective oxidation of hydrogen in order to preheat the incominghydrocarbon feedstock.

The subject process may be employed with a wide range of feedstocks. Thefeed hydrocarbon may therefore include C₆ -plus cyclic and acyclichydrocarbons including C₆ paraffins, C₇ paraffins, ethylbenzene andother alkylaromatic hydrocarbons. Other suitable feedstocks comprise alight hydrocarbon, a term used herein to refer to a hydrocarbon havingless than six carbon atoms per molecule including ethane. Theutilization of the subject process with any particular hydrocarbon willof course depend on an economic evaluation of the cost of utilizing thesubject process compared with the benefits with which it provides.

The conditions which will be employed in the dehydrogenation zone of theprocess will vary depending on such factors as feedstock, catalystactivity, and desired conversion. A general range of conditions whichmay be employed for dehydrogenation include a temperature from about550° C. to about 800° C., a pressure from about 0.3 to about 20atmospheres absolute and a liquid hourly space velocity from about 0.5to about 20 hr⁻¹. For the dehydrogenation of propane, for example, thepreferred conditions include a temperature in the range of about 600° C.to about 700° C., a pressure from about 1 to about 3 atmospheres, aliquid hourly space velocity of about 1 to about 10 hr⁻¹ and a hydrogento total hydrocarbon ratio between 1:1 and 5:1. It is especiallypreferred that the inlet temperature to each bed of propanedehydrogenation catalyst is between 650° C. and 690° C. The pressure inthe reaction zone employed within the process preferably differs only bythe incidental pressure drop which occurs as the reactants pass throughthe overall reaction system. The pressure maintained in the selectiveoxidation zone is therefore essentially the same as the pressure in thedehydrogenation zone. The inlet temperature of the bed of selectiveoxidation catalyst is preferably between 10° and 20° C. below thedesired inlet temperature of the dehydrogenation catalyst in thedehydrogenation zone. Larger temperature increases in the oxidation zonemay be possible and desired depending on the selectivity of theoxidation catalyst and the degree of hydrogen conversion desired.

The dehydrogenation zone may utilize any convenient system for thecatalytic dehydrogenation of the hydrocarbon feedstock. Thedehydrogenation zone preferably comprises at least one radial flowreactor in which the catalyst gradually moves downward by gravity flowto allow the near continuous replacement of used catalyst with catalysthaving a higher activity. Preferably, this replacement catalyst isregenerated in the appropriate facilities after being removed from thelowermost portion of the unitary multi-stage dehydrogenation reactionzone. It is preferred that a multi-stage reaction zone in which thereactants make at least two, preferably three, passes through a catalystbed is employed. The dehydrogenation catalyst therefore preferablyenters the top of a single unitary outer vessel containing the separatedehydrogenation stages and flows downward through the vessel from stageto stage by the action of gravity. A detailed description of moving bedreactors may be obtained by reference to U.S. Pat. Nos. 3,647,680;3,652,231; 3,706,536; 3,785,963; 3,825,116; 3,839,196; 3,839,197;3,854,887; and 3,856,662.

A preferred propane dehydrogenation catalyst comprises a platinum groupcomponent, a tin component and an alkali metal component and a porousinorganic carrier material. Other catalytic compositions may be usedwithin the dehydrogenation zone if desired.

It is preferred that the porous carrier material of the hereinabovedescribed dehydrogenation catalyst is an absorptive high surface areasupport having a surface area of about 25 to about 500 m² /g. The porouscarrier material should be relatively refractory to the conditionsutilized in the reaction zone and may be chosen from those carriermaterials which have traditionally been utilized in dualfunctionhydrocarbon conversion catalysts. A porous carrier material maytherefore be chosen from an activated carbon, coke or charcoal, silicaor silica gel, clays and silicates including those syntheticallyprepared and naturally occurring, which may or may not be acid-treated,as for example attapulgus clay, diatomaceous earth, kieselguhr, bauxite;refractory inorganic oxides such as alumina, titanium dioxide, zirconiumdioxides, magnesia, silica-alumina, alumina-boria; crystallinealuminosilicates such as naturally occurring or synthetically preparedmordenite or a combination of one or more of these materials. Apreferred porous carrier material is a refractory inorganic oxide, withthe best results being obtained with an alumina carrier material. Thecrystalline aluminas, such as gamma-alumina, give the best results. Ingeneral, the preferred catalyst will have a gamma-alumina carrier whichis in the form of spherical particles having a relatively small diameteron the order of about 1/16 inch.

A preferred alumina carrier material for the dehydrogenation catalystmay be prepared in any suitable manner. For example, the alumina carriermay be prepared by adding a suitable alkaline reagent, such as ammoniumhydroxide, to a salt of aluminum such as aluminum chloride in an amountto form an aluminum hydroxide gel which upon drying and calcining isconverted to alumina. It is particularly preferred that alumina spheresare manufactured by the well-known oil drop method which comprisesforming an alumina hydrosol by these techniques taught in the art, andpreferably by reacting aluminum metal with hydrochloric acid andcombining the hydrosol with a suitable gelling agent. The resultantmixture is dropped into an oil bath maintained at elevated temperatures.The droplets remain in the oil bath until they set and form hydrogelspheres. The spheres are then continuously withdrawn from the oil bathand are normally subjected to specific aging treatments in oil and anammoniacal solution to further improve their physical characteristics.The resulting pellets are then washed and dried at relatively lowtemperatures of about 150° C. to about 200° C. and calcined at atemperature of about 450° C. to about 700° C. for a period of about 1 toabout 20 hours. See the teachings of U.S. Pat. Nos. 2,620,314 and4,250,058 for additional details on the preparation of the base materialby the oil drop method.

A preferred dehydrogenation catalyst as described hereinabove alsocontains a platinum group component. Of the platinum group metals whichinclude palladium, rhodium, ruthenium, osmium, and iridium, the use ofplatinum is preferred. The platinum group components may exist withinthe final catalyst composite as a compound such as an oxide, sulfide,halide, oxysulfide, or as an elemental metal or in combination with oneor more other ingredients of the catalyst. It is believed that bestresults are obtained when substantially all the platinum group componentexists in the elemental state. The platinum group component generallycomprises from about 0.01 to about 2 weight percent of the finalcatalytic composite, calculated on an elemental basis. It is preferredthat the platinum content of the catalyst is between about 0.1 and 1weight percent. The preferred platinum group component is platinum, withpalladium being the next preferred metal. The platinum group componentmay be incorporated into the catalyst composite in any suitable mannersuch as by coprecipitation or cogelation with the preferred carriermaterial, or by ion-exchange or impregnation of the carrier material.The preferred method of preparing the catalyst normally involves theutilization of a water-soluble, decomposable compound of a platinumgroup metal to impregnate the calcined carrier material. For example,the platinum group component may be added to the support by comminglingthe support with an aqueous solution of chloroplatinic or chloropalladicacid. An acid such as hydrogen chloride is generally added to theimpregnation solution to aid in the distribution of the platinum groupcomponent throughout the carrier material.

The tin component of the hereinabove mentioned dehydrogenation catalystshould constitute about 0.01 to about 5 weight percent of the finalcomposite, calculated on an elemental basis, although substantiallyhigher amounts of tin may be utilized in some cases. Best results areoften obtained with about 0.1 to about 1 weight percent tin. It ispreferred that the atomic ratio of tin to platinum is between 1:1 and6:1. The tin component may be incorporated into the catalytic compositein any suitable manner known to effectively disperse this component in avery uniform manner throughout the carrier material. Thus, the componentmay be added to the carrier by coprecipitation. A preferred method ofincorporating the tin component involves coprecipitation during thepreparation of the preferred carrier material. This method typicallyinvolves the addition of a suitable soluble tin compound, such asstannous or stannic chloride to an alumina hydrosol, mixing theseingredients to obtain a uniform distribution throughout the sol and thencombining the hydrosol with a suitable gelling agent and dropping theresultant admixture into the oil bath as previously described. The tincomponent may also be added through the utilization of a soluble,decomposable compound of tin to impregnate the calcined porous carriermaterial. A more detailed description of the preparation of the carriermaterial and the addition of the platinum component and the tincomponent to the carrier material may be obtained by reference to U.S.Pat. No. 3,745,112.

A preferred catalyst as hereinabove described contains an alkali metalcomponent chosen from cesium, rubidium, potassium, sodium and lithium.The preferred alkali metal is normally either potassium or lithiumdepending on the feed hydrocarbon. The concentration of the alkali metalmay range from between 0.1 and 3.5 weight percent but is preferablybetween 0.2 and about 2.5 weight percent calculated on an elementalbasis. This component may be added to the catalyst by the methoddescribed above as a separate step or simultaneously with the additionof another component. With some alkali metals, it is normally necessaryto limit the halogen content to less than 0.5 weight percent andpreferably less than 0.1 weight percent.

Dehydrogenation catalysts generally consist of one or more metalliccomponents selected from Groups VI and VIII of the Periodic Table. Onetypical catalyst for the dehydrogenation of alkylaromatic hydrocarbonscomprises 85 percent by weight ferric oxide, 2 percent chromia, 12percent potassium hydroxide and 1 percent sodium hydroxide. A seconddehydrogenation catalyst, which is used commercially, consists of 87-90percent ferric oxide, 2-3 percent chromium oxide and 8-10 percentpotassium oxide. A third typical catalyst comprises 90 percent by weightiron oxide, 4 percent chromia and 6 percent potassium carbonate. Methodsfor preparing suitable catalysts are well known to the art. This isdemonstrated by the teachings of U.S. Pat. No. 3,387,053, whichdescribes the manufacture of a catalytic composite of at least 35 weightpercent iron oxide as an active catalytic agent, from about 1-8 weightpercent zinc or copper oxide, about 0.5-50 weight percent of an alkalipromoter and from about 1-5 weight percent chromic oxide as a stabilizerand a binding agent. U.S. Pat. No. 4,467,046 also describes a catalystfor the dehydrogenation of ethylbenzene in the presence of steam. Thiscatalyst consists of 15 to 30 weight percent potassium oxide, 2 to 8percent cerium oxide, 1.5 to 6 percent molybdenum oxide, 1 to 4 percentcalcium carbonate, with the balance iron oxide.

The selective oxidation catalyst employed in the subject process may beany commercially suitable catalyst which meets the required standardsfor stability and activity and which possesses high selectivity for theoxidation of hydrogen as compared to the oxidation of the feed orproduct hydrocarbon. That is, the oxidation catalyst must have a highselectivity for the oxidation of hydrogen with only small amounts of thefeed or product hydrocarbon being oxidized. A preferred oxidationcatalyst comprises a Group VIII noble metal, a Group IVA metal and ametal or metal cation which possesses a crystal ionic radius greaterthan 1.3 A, with these materials being present in small amounts on arefractory solid support. The preferred Group VIII metals are platinumand palladium, but the use of ruthenium, rhodium, osmium and iridium isalso contemplated. The Group VIII metal is preferably present in anamount equal to 0.01 to 5 weight percent of the finished catalyst. Ofthe metals of Group IVA of the Periodic Table, germanium, tin and leadcomprise the preferred species, these metals being present in the finalcatalyst composite in an amount from about 0.01% to about 5% by weight.The metal or metal cation having a radius greater than 1.3 A ispreferably chosen from Groups IA or IIA and is present in an amountequal to 0.01 to about 10 weight percent of the finished catalyst. Thiscomponent of the catalyst may include lithium, sodium, potassium,cesium, rubidium, calcium, francium, strontium and barium. Aparticularly preferred oxidation catalyst comprises platinum, tin andlithium. Further details of a preferred selective oxidation catalyst areprovided in U.S. Pat. No. 4,652,687 (Imai et al).

The preferred solid support for the selective oxidation catalyst isalumina having a surface area between 1 and 300 m² /g and apparent bulkdensity of between about 0.2 and 1.5 g/cc and an average pore sizegreater than 20 A. The metal-containing components are preferablyimpregnated into solid particles of the solid support by immersion in anaqueous solution followed by drying and calcination at a temperature offrom about 500° C. to 600° C. in air. The support may be in the form ofspheres, pellets or extrudates.

The operating conditions utilized during the contacting of thedehydrogenation zone effluent stream with the oxidation catalystcontained in the selective oxidation zone will be, to a large extent,set by the previously referred to dehydrogenation conditions. The inlettemperature of the selective oxidation zone will be necessarilydetermined by the temperature of the dehydrogenation zone effluent, andthe temperature of the oxygen which is introduced into the selectiveoxidation catalyst contained in the selective oxidation zone. Theincrease in temperature in the selective oxidation zone during thepassage of the reactant stream will depend upon the available quantityof hydrogen and the amount of oxygen introduced into the selectiveoxidation zone. The outlet temperature of the selective oxidation zonewill be necessarily influenced by the degree of selective oxidation ofthe hydrogen within the oxidation zone. In accordance with the presentinvention, a prime objective is to combust or oxidize a significantamount of the hydrogen present in the dehydrogenation zone effluentstream in the selective oxidation zone in order to reduce the quantityof hydrogen which must be subsequently compressed and to generate heatwhich may be recovered and utilized within the process and particularlyto heat the incoming fresh feed stream. The temperature and pressureconditions maintained in the selective oxidation zone preferably includea temperature from about 500° C. to about 800° C. and a pressure fromabout 0.3 to about 20 atmospheres absolute. The space velocity throughthe oxidation catalyst may preferably range from about 0.5 to about 20hr⁻¹. The liquid hourly space velocity, based on the liquid hydrocarboncharge at 60° F. (15° C.) is preferably between about 1 and about 20hr⁻¹. It is that substantially all of the oxygen which enters theselective oxidation zone is consumed within that zone and that theeffluent stream from the selective oxidation zone contains less than 0.1mole percent oxygen. The total amount of oxygen charged to the selectiveoxidation zone may be readily determined by an artisan based upon thedegree of hydrogen consumption desired and the outlet temperature of theselective oxidation zone which is required. The oxygen source may beair, but it is preferred that an oxygen-enriched gas containing lessthan 5 mole percent of nitrogen or other impurities is used as theoxygen source. To avoid any cooling of the selective oxidation zoneeffluent the oxygen-containing gas stream should preferably be heated toa temperature equal to the temperature of the inlet to the selectiveoxidation zone. In accordance with the present invention, the selectiveoxidation catalyst may be employed or installed in the selectiveoxidation zone in any convenient manner known in the art.

The effluent stream removed from the selective oxidation zone is heatexchanged for the purpose of lowering its temperature for the recoveryof heat. This effluent stream may be heat exchanged against a stream ofsteam, a reactant stream of this or another process or used as a heatsource for fractionation. Commercially, this effluent stream is oftenpassed through several heat exchanges thereby heating a number ofdifferent streams including the incoming hydrocarbon feedstock. Thisheat exchange is performed subject to the constraints set out above. Thesubsequently cooled effluent stream from the selective oxidation zone ispassed into a fractionation zone as a mixed phase stream to allow thefacile crude separation by decantation of the hydrocarbons from thewater and any net gas including any unreacted hydrogen present in theeffluent stream. In accordance with one embodiment of the presentinvention, the styrene present in the fractionation zone becomes part ofa hydrocarbon stream which is withdrawn from the fractionation zone andtransferred to proper downstream separation facilities. Preferably, thestyrene or other product hydrocarbon is recovered from the hydrocarbonstream by using one of the several fractionation systems known in theart. This fractionation will preferably yield a relatively pure streamof ethylbenzene, which is recycled, and an additional stream comprisingbenzene and toluene. These two aromatic hydrocarbons are by-products ofthe dehydrogenation reaction. They may be recycled in part as taught inU.S. Pat. No. 3,409,689 and British Pat. No. 1,238,602 or entirelyrejected from the process. Styrene is recovered as a third stream whichis withdrawn from the process. If desired, methods other thanfractionation may be used to recover the styrene. For instance, U.S.Pat. No. 3,784,620 teaches the separation of styrene and ethylbenzenethrough the use of a polyamide permeation membrane such as nylon-6 andnylon-6,10. U.S. Pat. No. 3,513,213 teaches a separatory methodemploying liquid-liquid extraction in which anhydrous silverfluoroborate is used as the solvent. Similar separatory methodsutilizing cuprous fluoroborates and cuprous fluorophosphates aredescribed in U.S. Pat. Nos. 3,517,079; 3,517,080 and 3,517,081.

The recovery of styrene through the use of fractionation is described inseveral references including U.S. Pat. No. 3,525,776. In this reference,the hydrocarbonaceous phase removed from the initial fractionation zoneis passed into a first column referred to as a benzene-toluene column.This column is operated at a subatmospheric pressure to allow itsoperation at lower temperatures and hence reduce the rate of styrenepolymerization. Various inhibitors such as elemental sulfur for exampleare injected into the column for this same purpose. Sulfur can also beintroduced into this column by returning at least a portion of the highmolecular weight material separated from the bottom stream of a styrenepurification column. A more detailed description of this is contained inU.S. Pat. Nos. 3,476,656; 3,408,263; and 3,398,063. There is effectedwithin the benzene-toluene column a separation of benzene and toluenefrom the effluent to produce an overhead stream which is substantiallyfree of styrene and ethylbenzene. This stream preferably contains atleast 95 mole percent benzene and toluene. The bottoms of thebenzene-toluene column is passed into second fractionation column fromwhich ethylbenzene is removed as an overhead product and recycled. Thebottoms stream of this column is then purified to obtain the styrene.Product recovery techniques directed to the recovery of vinyl toluenevia fractionation and the use of chemical additives to inhibitpolyerization are described in U.S. Pat. Nos. 4,417,085 and 4,492,675.The use of inhibitors and alternative fractionation techniques forreadily polymerizable vinyl aromatic compounds is also described in U.S.Pat. No. 4,469,558.

As previously mentioned, the subject process is not limited to theproduction of styrene and may be used to produce para methyl styrene bydehydrogenation of ethyl toluene or for the production of otherunsaturated product hydrocarbons such as acyclic C₃ -C₈ olefins. Theproduct hydrocarbon recovered from the process may therefore bepropylene, a butylene or a mixture of butylenes, a heptene, etc.

The foregoing description and drawing clearly illustrate the advantagesencompassed by the process of the present invention and the benefits tobe afforded with the use thereof.

We claim as our invention:
 1. A process for the catalyticdehydrogenation of a dehydrogenatable C₂ -plus feed hydrocarbon whichcomprises the steps of:(a) passing a feed stream comprising said C₂-plus feed hydrocarbon into a dehydrogenation zone and through at leastone bed of dehydrogenation catalyst maintained at dehydrogenationconditions and producing a dehydrogenation zone effluent streamcomprising hydrogen, the C₂ -plus feed hydrocarbon and a C₂ -plusproduct hydrocarbon; (b) forming an oxidation catalyst bed feed streamby admixing an oxygen-containing stream with said dehydrogenation zoneeffluent stream; (c) passing the oxidation catalyst bed feed streamthrough a bed of hydrogen selective oxidation catalyst maintained atselective oxidation conditions and producing an oxidation zone effluentstream having a reduced concentration of hydrogen; and (d) recoveringthe product hydrocarbon from said oxidation zone effluent stream withoutcontacting said oxidation zone effluent stream with dehydrogenationcatalyst.
 2. The process of claim 1 wherein said feed hydrocarboncomprises alkylaromatic hydrocarbons.
 3. The process of claim 2 whereinsaid alkylaromatic hydrocarbons are the group consisting ofethylbenzene, diethylbenzene, ethyl toluene, propyl benzene andisopropyl benzene.
 4. The process of claim 1 wherein said feedhydrocarbon comprises C₂ -C₁₆ paraffins.
 5. The process of claim 4wherein said C₂ -C₁₆ paraffins are selected from the group consisting ofpropane, butane, hexane and nonane.
 6. The process of claim 1 whereinsaid dehydrogenation catalyst comprises a platinum group component, atin component, an alkali metal component and a porous inorganic carriermaterial.
 7. The process of claim 1 wherein said dehydrogenationcatalyst comprises ferric oxide and chromia.
 8. The process of claim 1wherein said dehydrogenation conditions include a temperature from about550° C. to about 800° C., a pressure from about 0.3 to about 20atmospheres absolute and a liquid hourly space velocity from about 0.5to about 20 hr⁻¹.
 9. The process of claim 1 wherein said hydrogenselective oxidation catalyst comprises a Group VIII noble metal and ametal or metal cation which possesses a crystal ionic radius greaterthan 1.3 A.
 10. The process of claim 1 wherein said hydrogen selectiveoxidation catalyst comprises platinum, lithium and alumina.
 11. Theprocess of claim 1 wherein said selective oxidation conditions include atemperature from about 500° C. to about 800° C., a pressure from about0.3 to about 20 atmospheres absolute and a liquid hourly space velocityfrom about 1 to about 20 hr⁻¹.
 12. The process of claim 1 wherein saidoxygen-containing stream contains less than 5 mole percent of nitrogen.13. A process for the catalytic dehydrogenation of a dehydrogenatable C₂-plus feed hydrocarbon which comprises the steps of:(a) passing a feedstream comprising said C₂ -plus feed hydrocarbon into a dehydrogenationzone and through at least one bed of dehydrogenation catalyst maintainedat dehydrogenation conditions and producing a dehydrogenation zoneeffluent stream comprising hydrogen, the C₂ -plus feed hydrocarbon and aC₂ -plus product hydrocarbon; (b) forming an oxidation catalyst bed feedstream by admixing an oxygen-containing stream with said dehydrogenationzone effluent stream; (c) passing the oxidation catalyst bed feed streamthrough a bed of hydrogen selective oxidation catalyst maintained atselective oxidation conditions and producing an oxidation zone effluentstream having a reduced concentration of hydrogen; (d) heat exchangingsaid oxidation zone effluent stream with said feed stream in step (a) toraise the temperature of said feed stream and to lower the temperatureof said oxidation zone effluent stream; and (e) recovering the producthydrocarbon from said oxidation zone effluent stream from step (d)without contacting the oxidation zone effluent stream withdehydrogenation catalyst.
 14. The process of claim 13 wherein said feedhydrocarbon comprises alkylaromatic hydrocarbons.
 15. The process ofclaim 14 wherein said alkylaromatic hydrocarbons are selected from thegroup consisting of ethylbenzene, diethylbenzene, ethyl toluene, propylbenzene and isopropyl benzene.
 16. The process of claim 13 wherein saidfeed hydrocarbon comprises C₂ -C₁₆ paraffins.
 17. The process of claim15 wherein said C₂ -C₁₆ paraffins are selected from the group consistingof propane, butane, hexane and nonane.
 18. The process of claim 13wherein said dehydrogenation catalyst comprises a platinum groupcomponent, a tin component, an alkali metal component and a porousinorganic carrier material.
 19. The process of claim 13 wherein saiddehydrogenation catalyst comprises ferric oxide and chromia.
 20. Theprocess of claim 13 wherein said dehydrogenation conditions include atemperature from about 550° C. to about 800° C., a pressure from about0.3 to about 20 atmospheres absolute and a liquid hourly space velocityfrom about 0.5 to about 20 hr⁻¹.
 21. The process of claim 13 whereinsaid hydrogen selective oxidation catalyst comprises a Group VIII noblemetal and a metal or metal cation which possesses a crystal ionic radiusgreater than 1.3 A.
 22. The process of claim 13 wherein said hydrogenselective oxidation catalyst comprises platinum, lithium and alumina.23. The process of claim 13 wherein said selective oxidation conditionsinclude a temperature from about 500° C. to about 800° C., a pressurefrom about 0.3 to about 20 atmospheres absolute and a liquid hourlyspace velocity from about 1 to about 20 hr⁻¹.
 24. The process of claim13 wherein said oxygen-containing stream contains less than 5 molepercent of nitrogen.